Process for the polymerization of olefins

ABSTRACT

The present invention relates to a process for the continuous preparation of polypropylene in a reactor from propylene and optionally ethylene and/or at least one other oolefin monomer, wherein the reactor comprises a fluidized bed, an expanded section located at or near the top of the reactor, a distribution plate located at the lower part of the reactor and an inlet for a recycle stream located under the distribution plate, wherein the process comprises—feeding a polymerization catalyst to the fluidized bed in the area above the distribution plate—feeding the propylene and the optional at least one other oolefin monomer to the reactor—withdrawing the polypropylene from the reactor—circulating fluids from the top of the reactor to the bottom of the reactor, wherein the circulating fluids are compressed using a compressor and cooled using a heat exchanger, resulting in a cooled recycle stream comprising liquid, and wherein the cooled recycle stream is introduced into the reactor using the inlet for the recycle stream wherein an alkane chosen from the group of iso-butane, n-butane, cyclopropane and mixtures thereof is added to the reactor and wherein the molar composition of the components in the recycle stream is chosen such that the dew temperature of the recycle stream at the reactor pressure is at least 0.10° C. below the temperature of the reactor.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a 371 of International Application No.PCT/EP2016/067874, filed Jul. 27, 2016, which claims priority toEuropean Patent No. 15180261.8, filed Aug. 7, 2015 and U.S. ApplicationNo. 62/261,502 filed Dec. 1, 2015, all of which are incorporated hereinby reference in their entirety.

The present invention relates to a process for the preparation of apolypropylene in a reactor from one or more α-olefin monomers of whichat least one is ethylene or propylene. The invention also relates topolyolefins obtained or obtainable by said process and to a reactionsystem for operating said process.

Processes for the homopolymerization and copolymerization of olefins forexample to produce homopolypropylene or a propylene-ethylene copolymerin the gas phase are well known in the art. Typically, in fluidized bedpolymerization process, the solid particulates are projected upward intothe expanded section through the bursting of rising gas bubbles at thesurface of the fluidized bed, and most of these particulates aretypically returned to the fluidized bed by gravity as their velocitydissipates in the lower gas velocities of the expanded section. A smallquantity of fine powder, or fines, is elutriated out of the projectedparticulates and does not return directly to the fluid bed by gravity.These fines are either conveyed upward by the cycle gas to the reactorgas outlet, or settle on surfaces of the expanded section throughgravity or through particle attraction forces such as electrostaticattraction.

Disengaged fines that settle on the expanded section surfaces are knownto accumulate as layers of fines under certain conditions. Settled finesare generally reactive and continue to polymerize in place at a raterelated to the concentration of active catalyst contained in the fines.Such layers typically build to sufficient thicknesses in a short periodof time that the forces holding them in place are overcome by gravityand the layers then slide harmlessly back into the fluid bed. Largerparticulates from the fluid bed may also be projected onto the layers offines, especially at lower elevations near the bed surface, causing allor part of the layer to be released and to then re-enter the bed throughgravity. The cycle of fines buildup and return to the bed occursrepetitively in normal operation.

It is well known that the production rate (i.e. the space time yield interms of weight of polymer produced per unit volume of reactor space perunit time) in commercial gas fluidized bed reactors of theafore-mentioned type is restricted by the maximum rate at which heat canbe removed from the reactor. The rate of heat removal can be increasedfor example, by increasing the velocity of the recycle gas and/orreducing the temperature of the recycle gas. However, there is a limitto the velocity of the recycle gas which can be used in commercialpractice.

Beyond this limit the bed can become unstable or even lift out of thereactor in the gas stream, leading to blockage of the recycle line anddamage to the recycle gas compressor or blower. There is also a limit onthe extent to which the recycle gas can be cooled in practice. This isprimarily determined by economic considerations, and in practice isnormally determined by the temperature of the industrial cooling wateravailable on site. Refrigeration can be employed if desired, but thisadds to the production costs. Thus, in commercial practice, the use ofcooled recycle gas as the sole means of removing the heat ofpolymerization from the gas fluidized bed polymerization of olefins hasthe disadvantage of limiting the maximum production rates obtainable.

Conventionally, gas phase polymerization processes typically runcontinuously, therefore the temperature of the fluidized bed reactor iscontrolled to an essentially isothermal level through continuouslyremoving the heat of polymerization by circulating the gas exiting fromthe fluidized bed to a condenser/heat exchanger outside the reactor andrecirculating the cooled gas stream back into the reactor. When thetemperature of the recirculating stream introduced or recycled into thefluidized bed polymerization reactor is above the dew point temperature,substantially no liquid is present. This process is known as the “drymode” process. One method to maximize the ability of heat removal is,throughout the operation, to reduce to the lowest possible value thetemperature of the gaseous feed stream into the reactor.

The prior art suggests a number of methods for removing heat from gasfluidized bed polymerization processes. According to the “condensedmode” process a two phase mixture is directed into the fluidized bed asa fluidizing medium, the liquid portion of which vaporizes when it isexposed to the heat of the reactor. Fluid is formed by cooling therecycle stream below the dew point temperature, thereby converting aportion of the gas into a liquid, and the cooled recycle stream isintroduced into the fluidized bed polymerization reactor. The objectiveis to take advantage of the cooling effect brought about by thevaporization, i.e., by bringing the temperature of the fluidized beddown to a point where degradation of the polymer and the catalyst areavoided and agglomeration of the polymer and chunking are prevented. Theliquid phase is provided by a portion of the recycle gases, whichincludes monomers and low boiling liquid hydrocarbons, inert to thereaction conditions needed for polymerization, and condensation.Condensed mode fluidized bed reactor polymerization processes have beendisclosed in for example U.S. Pat. No. 4,543,399 and 4,588,790 whichdescribes introducing an inert liquid into the recycle stream toincrease the dew point temperature of the recycle stream and allowingthe process to operate at levels of up to 17.4% liquid by weight, basedon the total weight of the cooled recycle stream. A condensed modeprocess is considered to be advantageous because its ability to removegreater quantities of heat generated by polymerization increases thepolymer production capacity of a fluidized bed polymerization reactor. Acommonly used liquid hydrocarbon is isopentane, which boils at about 27°C. C., and consequently becomes a vapor in the recycle line in view ofthe heat present in the recycle gases. The recycle gases leave thereactor, are cooled, and then condensed to the extent that a vapor phaseand liquid phase are formed. The velocity of the recycled gas/liquidmixture should be sufficient to support the fluidized bed, but slowenough to avoid excessive entrainment of fines. The cooling capacityshould be sufficient to improve the production rate in terms ofspace/time/yield.

US2005/0182207 discloses a continuous gas fluidized bed polymerizationprocess for the production of a polymer from a monomer comprising:

-   continuously passing a gaseous stream comprising the monomer through    a fluidized bed reactor in the presence of a catalyst under reactive    conditions;-   withdrawing a polymeric product and a stream comprising unreacted    monomer gases;-   cooling said stream comprising unreacted monomer gases to form a    mixture comprising a gas phase and a liquid phase and reintroducing    said mixture into said reactor with sufficient additional monomer to    replace that monomer polymerized and withdrawn as the product,    wherein said liquid phase is vaporized, and wherein the stream    comprises at least two inert condensing agents selected from the    group consisting of alkanes, cycloalkanes, and mixtures thereof,    each of the inert condensing agents having a normal boiling point    less than 40° C.

However, this process has the disadvantage that particles (fines) arestill built up and lumps are formed. Fines and lump formation is a majorproblem in gas-phase polymerization processes as the more fines areproduced, the more down time a reactor will have due to the necessityfor maintenance.

Therefore, it is the object of the invention to provide a process forthe continuous preparation of a polypropylene in a reactor, whichreactor is operated in a condensed mode, wherein particle build up andlump formation is reduced.

This object is achieved by a process for the continuous preparation ofpolypropylene in a reactor from propylene and optionally at least oneother α-olefin monomer, wherein the reactor comprises a fluidized bed,an expanded section located at or near the top of the reactor, adistribution plate located at the lower part of the reactor and an inletfor a recycle stream located under the distribution plate

wherein the process comprises

-   -   feeding a polymerization catalyst to the fluidized bed in the        area above the distribution plate    -   feeding the propylene and the optional at least one other        α-olefin monomer to the reactor    -   withdrawing the polypropylene from the reactor    -   circulating fluids from the top of the reactor to the bottom of        the reactor, wherein the circulating fluids are compressed using        a compressor and cooled using a heat exchanger, resulting in a        cooled recycle stream comprising liquid, and wherein the cooled        recycle stream is introduced into the reactor using the inlet        for the recycle stream

wherein an alkane chosen from the group of iso-butane, n-butane,cyclopropane and mixtures thereof is added to the reactor and

wherein the molar composition of the components in the recycle stream ischosen such that the dew temperature of the recycle stream at thereactor pressure is at least 0.1° C. below the temperature of thereactor

It has been found that by adding an alkane chosen from the group ofiso-butane, n-butane, cyclopropane and mixtures thereof to the reactorand by choosing the molar composition of the components in the recyclestream chosen such that the

dew temperature of the recycle stream at the reactor pressure is atleast 0.1° C. below the temperature of the reactor, has the advantagethat the space time yield is increased while maintaining a stableprocess. In addition, agglomeration tendency may be decreased due toless stickiness of the particles produced as compared to other attemptsto increase the space time yield.

Additionally, the recovery of the hydrocarbons through the polymerdischarge system may be higher (due to the presence of higher amounts oflight hydrocarbons) while it is not necessary to increase the degassingefficiency for the polymer in the degassing section.

Also, the potential of forming dead zones (also known as hot spots) inthe fluidized bed may be reduced on local microscopic level. This willreduce the formation of fines and will increase the overall catalystproductivity for a catalyst which has a typical catalyst decay profile.

US2005/0137364 discloses a continuous gas fluidized bed polymerizationprocess for the production of a polymer from a monomer includingcontinuously passing a gaseous stream comprising the monomer through afluidized bed reactor in the presence of a catalyst under reactiveconditions; withdrawing a polymeric product and a stream comprisingunreacted monomer gases; cooling said stream comprising unreactedmonomer to form a mixture comprising a gas phase and a liquid phase andreintroducing said mixture into said reactor with sufficient additionalmonomer to replace that monomer polymerized and withdrawn as theproduct, wherein said liquid phase is vaporized, and wherein the streamcomprises an induced condensing agent selected from the group consistingof alkanes, cycloalkanes, and mixtures thereof, the induced condensingagent having a normal boiling point less than 25° C. However, theexamples are all related to the production of linear low densitypolyethylene (LLDPE) and therefore the teaching of this document is notapplicable to polypropylene production.

With ‘continuous polymerization of one or more α-olefins’ or ‘continuouspreparation of polyolefin’ is meant herein that one or more α-olefinmonomers of which at least one is ethylene or propylene are fed to thereactor and polyolefin thus produced is (semi)-continuously withdrawnthrough a polymer discharge system connected to the reactor.

Preferred other α-olefin monomers include for example α-olefins havingfrom 4 to 8 carbon atoms. However, small quantities of α-olefin monomershaving more than 8 carbon atoms, for example 9 to 18 carbon atoms, suchas for example a conjugated diene, can be employed if desired. Thus itis possible to produce homopolymers of ethylene or propylene orcopolymers of ethylene and/or propylene with one of more α-olefinmonomers having from 4 to 8 α-olefin monomers. Preferred α-olefinmonomers include but are not limited to but-1-ene, isobutene,pent-1-ene, hex-1-ene, hexadiene, isoprene, styrene, 4-methylpent-1-ene,oct-1-ene and butadiene. Examples of α-olefin monomers having more than8 carbon atoms that can be copolymerized with an ethylene and/orpropylene monomer, or that can be used as partial replacement forα-olefin monomers having from 4 to 8 α-olefin monomers include but arenot limited to dec-1-ene and ethylidene norbornene.

When the system or process of the invention is used for thecopolymerization of ethylene and/or propylene with α-olefin monomers,the ethylene and/or propylene preferably is used as the major componentof the copolymer. For example, the amount of ethylene and/or propylenepresent in the copolymer is at least 65% by weight, for example at least70% by weight, for example at least 80% by weight based on the totalcopolymer.

A reactor is herein meant a vessel designed for reactions to take placetherein, comprising inlets for receiving feed materials and outlets fordischarging reaction products.

The reactor in the process of the invention comprises a fluidized bed,an expanded section located at or near the top of the reactor, adistribution plate located at the lower part of the reactor and an inletfor a recycle stream located under the distribution plate. The reactoris preferably closed off at the top and the bottom by a hemisphere.

With ‘fluidized bed’ as used herein is meant that particles (in thiscase preferably the catalyst and/or the catalyst to which one or moreα-olefin monomers of which at least one is ethylene or propylene isattached, herein also referred to as prepolymerized catalyst, and thegrowing polymer particles) are held suspended by the fluid stream inwhich a particle/fluid mixture acts as a fluid. This can be achieved byplacing the amount of particles under appropriate conditions, forinstance by upwardly flowing fluid (gas and or gas/liquid mixture)through the solid particles at a velocity that exceeds the minimumfluidization velocity and enough to suspend the solid particles andcausing them to behave as a fluid. The velocity should not be so high asto result in undue removal of polymer particles from the bed.

In order to maintain a fluidized bed in the process of the invention,the superficial gas velocity m may be in the range of 0.5 to 5 m/s. Forexample, may be at least 1, for example at least 1.5, for example atleast 2 and/or for example at most 4.5, for example at most 4, forexample at most 3.5, for example at most 3 m/s.

The expanded section located at or near the top of the reactor is notintended for gas-phase polymerization, but instead is suitable for gasexpansion. It has the function to disengage the reaction mixture and thepolymer product of the reaction. Accordingly, this section does notfunction as a reaction zone. The superficial gas velocity may be of suchlow value that polymer particles preferably do not enter into theexpanded section for example to avoid clogging to occur in thecompressor.

In such reactor, during the course of polymerization, fresh polymerparticles are produced by catalytic polymerization of α-olefin monomers.Such polymer particles are projected upwards in the direction of theexpanded section through the circulating gas. Most of these particles dopreferably not reach the expanded section or return to the fluidized bedby gravity as the superficial gas velocity decreases in the expandedsection.

The distribution plate may be any device that is suitable fordistributing the cooled recycle stream in the reactor to keep afluidized bed and to serve as a support for a quiescent bed of thepolymerization catalyst and polyolefin when the reactor is not inoperation. The distribution plate is used for achieving good gasdistribution. For example, the distribution plate may be a screen,slotted plate, perforated plate, a plate of the bubble-cap type, orother conventional or commercially available plate or other fluiddistribution device. An example of a commonly used type of distributionplate is a perforated plate with some above-hole structure on top ofeach hole, to prevent particle sifting.

The distribution plate is generally positioned perpendicular to thelongitudinal axis of a reactor, wherein the fluidized bed is locatedabove said distribution plate and a mixing chamber region below saiddistribution plate.

In addition to the distribution plate, the reactor may be furtherequipped with other means for agitation, such as mechanical agitation,for example a stirrer. Preferably, the reactor does not comprisemechanical agitation.

The person skilled in the art is aware of the type of inlets that may besuitable for allowing the recycle stream to enter into the reactor underthe distribution plate.

Catalyst

With ‘catalyst’ as used herein is meant to include both catalyst andcocatalyst, and any other compounds, which assist in catalyzing theproduction of the polyolefin as well as the prepolymerized catalyst (thecatalyst to which one or more α-olefin monomers of which at least one isethylene or propylene is attached).

The polymerization catalyst may be fed to the reactor for example byusing feeding means, such as a pump. The polymerization catalyst may forexample be fed as a suspension in a solvent, for example a hydrocarbonsolvent or the like, or in an inert gas, such as nitrogen (drycatalyst). The polymerization catalyst may also be injected into thefluidized bed.

The polymerization catalyst may be fed at any position in the area abovethe distribution plate or at a combination of positions in the reactor.

The person skilled in the art is aware of which catalysts are suitablefor continuous polymerization of monomers such as α-olefin monomers.

For example, heterogeneous polymerization catalysts, which are catalyststhat are supported on an inert substrate, for example silica or aluminamay be used. Suitable examples of heterogeneous catalysts includesupported Ziegler Natta and supported metallocene catalysts andcombinations thereof, for example in a mixed catalyst system. Examplesof a catalyst composition for polymerization of α-olefins comprising atleast two catalytic components are for example described in EP1764378A1,hereby incorporated by reference. EP1764378A1 discloses a catalystcomposition comprising a metallocene component and a Ziegler-Natta typetransition metal component, at least one activator and support material.Metallocene catalysts are for example described by Hamielec and Soaresin “Polymerisation reaction engineering-metallocene catalysts” (Prog.Pol. Sci. Vol. 21, 651-706, 1996), hereby incorporated by reference. Thepolymerization catalyst may also be a metal oxide catalyst, for examplea chromium oxide catalysts. Such metal oxide catalyst may for example bebased on a support of an inert substrate, for example on silica, aluminasilicate or alumina, for example on a highly porous support of silica,alumina silicate or alumina as for example disclosed in the “Handbook ofPolyethylene” by Andrew Peacock at pages 61-62, hereby incorporated byreference.

The group of metallocene catalysts includes many variations. In the mostgeneral form, metallocene catalysts comprise a metal atom, for exampletitanium, zirconium or hafnium attached to for example four ligands, forexample two substituted cyclopentadienyl ligands and two alkyl, halideor other ligands with an optionally modified organoalumoxane asactivator, for example methylaluminoxane (MAO) or a compound based onboron. Examples of inert substrates that can be used as support for ametallocene catalyst include inorganic oxides, for example SiO₂, MgCl₂,Al₂O₃, MgF₂ and CaF₂. Preferably, the polymerization catalyst used inthe system and process of the invention is a metallocene catalystsupported on silica, for example a silica as commercially available, forexample Grace Davison 948 silica or Ineos ES 70 silica.

A Ziegler Natta catalyst may be used together with a cocatalyst in thesystem and process of the invention. Suitable example of cocatalystsinclude but are not limited to organo aluminium compounds having formulaAIR₃, wherein R stands for a hydrocarbon having 1 to 10 C-atoms.Examples of organo aluminium compounds having formula AIR₃ includetriethylaluminium alkyl, triisobutyl aluminium trialkyl, tri-n-hexylaluminium and tri octyl aluminium.

Examples of inert substrates that can be used as support for a ZieglerNatta catalyst include inorganic oxides, for example oxides of silica,alumina, magnesium, titanium and/or zirconium; magnesium chloride,clays, zeolites, polystyrene, polyethylene, polypropylene, graphiteand/or layered silicates.

It will be clear to the person skilled in the art that also mixtures ofpolymerization catalysts may be used in the invention.

The optimal amount of catalyst and ratios of procatalyst to cocatalystand potential external donors can easily be determined by the personskilled in the art.

Ziegler-Natta catalyst systems and their components are suitable forpreparing a polyolefin and are generally known. “Ziegler-Natta” (Z-N)refers to catalyst systems comprising a transition metal-containingsolid catalyst compound (also typically referred to as a procatalyst);an organometallic compound (also typically referred to as a co-catalyst)and one or more electron donor compounds (e.g. external electrondonors). An overview of such catalyst types is for example given by T.Pullukat and R. Hoff in Catal. Rev.—Sci. Eng. 41, vol. 3 and 4, 389-438,1999. The preparation of such a procatalyst is for example disclosed inWO96/32427 A1. Several examples of Ziegler-Natta catalyst are discussedbelow.

EP 1 273 595 of Borealis Technology discloses a catalyst obtained byreacting a complex of a Gp IIa metal, an electron donor, and atransition metal as an emulsion the dispersed phase of which containsmore than 50 mol % of the Gp IIa metal in said complex.

EP 0 019 330 of Dow discloses a Ziegler-Natta type catalyst obtained byreacting a halogenated magnesium compound, a halide of tetravalenttitanium in the presence of a halohydrocarbon, and a tetravalenttitinanium compound.

The Examples of U.S. Pat. No. 5,093,415 of Dow discloses a catalystobtained by reacting titanium tetrachloride, diisobutyl phthalate, andmagnesium diethoxide, and then titanium tetrachloride and phthaloylchloride amd again titanium tetrachloride.

Example 2 of U.S. Pat. No. 6,825,146 of Dow discloses a catalystobtained by reacting titanium tetrachloride, magnesium diethoxide,titanium tetraethoxide, and titanium tetrachloride, with ethylbenzoate,and again titanium tetrachloride, and benzoylchloride.

U.S. Pat. No. 4,771,024 of Dow discloses a catalyst obtained by reactingdried silica with carbonated magnesium solution, then titaniumtetrachloride, chlorobenzene and ethylbenzoate, and TiCl₄ andchlorobenzene.

WO03/068828 of China Petroleum discloses a catalyst obtained by reactingmagnesium chloride, toluene, epoxy chloropropane and tributyl phosphate,then phthalic anhydride nd TiCl₄ and an internal donor, then againTiCl₄.

U.S. Pat. No. 4,866,022 of Amoco discloses a catalyst obtained byreacting a magnesium-containing species, a transition metal halide andan organosilane, again with transition metal compound and an electrondonor.

For example, the Ziegler-Natta type procatalyst in the catalyst systemaccording to the present invention is a magnesium-based supportedcatalyst obtained by the process as described in WO 2007/134851 A1,comprises the following phases: phase A): preparing a solid support forthe procatalyst based on a Grignard compound and a silane compound;phase B): optionally activating said solid support obtained in phase A)using one or more activating compounds to obtain an activated solidsupport; phase C): contacting said solid support obtained in phase A) orsaid activated solid support in phase B) with a titanium catalyticspecies and optionally one or more internal donors and/or optionally anactivator to obtain said procatalyst; optionally Phase D: modifying saidintermediate product obtained in phase C) with a Group 13- or transitionmetal modifier and optionally one or more internal donors.

The one or more α-olefin monomers may be fed to the reactor (8) usingfeeding means such as a pump. The monomers are preferably fed to thereactor (8) by adding the monomers to the fluids that are circulatedfrom the top of the reactor to the inlet for the recycle stream prior tocooling of the fluids. Preferably, the one or more α-olefin monomers areadded in such amounts that they make up for the one or more α-olefinmonomer consumed during the polymerization.

The one or more α-olefin monomers may be fed in one or in multiplefeeding streams. For example, one type of olefin monomer, typicallyethylene and/or propylene may be comprised in the feed (60) and anothertype of α-olefin monomer, also referred to herein as the comonomer, maybe comprised in the feed (70).

Likewise, withdrawal of the polyolefin (30) may be done at any positionin the area above the distribution plate or at a combination ofpositions.

Polyolefin (30) may be withdrawn from the multi-zone reactor (8) usingany suitable means, for example a polymer discharge system. Thepolyolefin may be used as such or may be subjected to purification orother end-processing.

Fluids are circulated from the top of the reactor to the bottom of thereactor. The circulating fluids are cooled using a heat exchanger,resulting in a cooled recycle stream comprising liquid. The cooledrecycle stream is introduced into the reactor using the inlet for therecycle stream.

For the avoidance of doubt the term ‘fluids’ encompasses liquids, gasesand mixtures thereof, wherein the term ‘liquids’ includes liquidscontaining solid particles, such as slurries.

The fluids may be cooled to below the dew point of the fluids using anysuitable cooling means. For example, cooling of the fluids may beperformed using a cooling unit, for example a heat exchanger.

The dew point may be increased by increasing the operating pressure ofthe fluids and/or by increasing the percentage of condensable fluids andsimultaneously decreasing the percentage of non-condensable gases in thefluids.

By feeding the fluids that are cooled to below the dew point of thefluids into the bottom inlet of the reactor, the fluids will be passedthrough the distribution plate (6) into the section above thedistribution plate (6), resulting in the formation of a fluidized bedand/or a bubble column. Heat generated by the polymerization will causethe liquids in the fluids to evaporate. The feeding of thepolymerization catalyst and the one or more α-olefin monomers to thereactor (8) will result in the formation of polyolefin (30), which iswithdrawn from the reactor (8). The recycle stream is recirculated fromthe top of the reactor to the bottom inlet. The one or more α-olefinmonomers and other fluids, such as hydrogen, an inert gas or liquid, forexample a condensable inert component, may be added to the recyclestream to make up for the target composition (there is loss due to e.g.reaction, withdrawal and venting) before cooling the fluids to below thedew point of the fluids to form a cooled recycle stream.

Preferably in the processes of the invention, the fluids are cooled tosuch extent that the amount of liquid in the cooled recycle stream (10)is at least 7% by weight, for example at least 9%, for example at least14% by weight based on the total amount of liquid and gas. For example,the amount of liquid in the cooled recycle stream is at least 14.5%, forexample at least 20%, for example at least 25% and/or for example atmost 95%, for example at most 90%, for example at most 90%, for exampleat most 85%, for example at most 80%, for example at most 75%, forexample at most 70%, for example at most 65%, for example at most 60%,for example at most 55%, for example at most 55% by weight based on thetotal amount of liquid and gas in the cooled recycle stream. Preferably,the amount of liquid in the cooled recycle stream is at least 25% andfor example at most 55% by weight based on the total amount of liquidand gas in said cooled recycle stream.

High amounts of liquid in the cooled recycle stream enables feeding ofone or more very high activity catalyst system.

The compressor (400) may be any device that is suitable for compressingthe feed (60) and the recycle stream (40) using the compressor (400) toform the compressed fluids (50). By compressing the feed (60) and therecycle stream (40), the pressure of the compressed fluids (50) isincreased compared to the feed (60) and the recycle stream (40) beforeuse of the compressor (400).

Introduction of the cooled recycle stream under the distribution platemay be done using any suitable means for introducing fluids, for exampleusing injection nozzles.

Preferably, wherein the amount of iso-butane, n-butane, cyclopropane andmixtures thereof is chosen such that the molar composition in therecycle stream of the total of iso-butane, n-butane and/or cyclopropaneis least 1 mol %, preferably at least 2 mol %, more preferably at least2.4 mol % and/or at most 15%, preferably at most 10 mol %, morepreferably at most 5 mol %. Molar composition of the reactor as definedherein is measured by measuring the amount of components in moles in therecycle stream upstream of the compressor.

Preferably, cyclopropane, iso-butane and/or n-butane are present in therecycle stream to replace propylene and/or alpha-olefin monomer (and donot substitute nitrogen) to not increase the dew point at the reactortemperature to above the reactor temperature.

In the process of the invention, the recycle stream comprisesiso-butane, n-butane and/or cyclopropane as well as propylene, propaneand nitrogen.

For example, the recycle stream comprises

-   1-10 mol %, for example 2-5 mol % of iso-butane, n-butane and/or    cyclopropane-   20-99.9 mol % propylene, for example 75-90 mol % propylene-   8-10 mol % propane-   0-3 mol % hydrogen, for example 0-2, for example 0.01 to 0.05 mol %    hydrogen-   0-10 mol % nitrogen, for example from 5-10 mol % nitrogen

wherein the total amount of mol % in the recycle stream is 100%.

Preferably, the alkane chosen from the group of iso-butane, n-butane,cyclopropane and mixtures thereof, is iso-butane or n-butane, morepreferably iso-butane. Preferably, the alkane in the alkane chosen fromthe group of iso-butane, n-butane, cyclopropane and mixtures comprisesat least 80 mol % iso-butane, more preferably at least 90 mol %iso-butane, even more preferably the alkane in the alkane chosen fromthe group of iso-butane, n-butane, cyclopropane and mixtures thereof isiso-butane.

For example, iso-butane is present in the recycle stream in not morethan 2.65 mol % (as a higher molar composition may lead to lesssolubility, may have an effect on the dew point and may lead tostickiness of the polypropylene). For the same reasons as foriso-butane, preferably, n-butane is for example present in the recyclestream in not more than 2.3 mol % and cyclopropane is present in therecycle stream in not more than 3.9 mol %.

The dew temperature of the recycle stream at the reactor pressure is atleast 0.10° C. below the temperature of the reactor. For example, thedew temperature of the recycle stream at the reactor pressure is atleast 0.11, for example at least 0.12, for example at least 0.13, forexample at least 0.14, for example at least 0.15, for example at least0.16, for example at least 0.17, for example at least 0.18, for exampleat least 0.19, for example at least 0.20° C. below the temperature ofthe reactor.

The alkane chosen from the group of iso-butane, n-butane, cyclopropaneand mixtures thereof may be added to the recycle stream by addition ofthe alkane somewhere to the reaction system, wherein the reaction systemcomprises the reactor and the recycle stream.

For example, the alkane may be added to the reactor under thedistribution plate or may be added to the expanded section.

In one embodiment, the alkane chosen from the group of iso-butane,n-butane, cyclopropane and mixtures thereof is introduced to the recyclestream downstream of the compressor and upstream of the heat exchanger.

In another embodiment, a liquid-containing stream comprising the alkanechosen from the group of iso-butane, n-butane, cyclopropane and mixturesthereof is introduced into the expanded section during at least part ofthe polymerization process, wherein this liquid-containing stream ispreferably brought into contact with at least part of the interiorsurface of the expanded section.

In the process of the invention, a liquid-containing stream isintroduced into the expanded section during at least part of thepolymerization process, for example the liquid-containing stream isintroduced into the expanded section during the entire polymerizationprocess. The addition of a liquid-containing stream may be intermittentor continuous.

For example, in this embodiment, the liquid-containing stream comprisingthe alkane chosen from the group of iso-butane, n-butane, cyclopropaneand mixtures thereof may be added to the expanded section in acircumferential manner, for example by using at least one nozzle that istangential to the interior wall of the expanded section or the liquidcontaining stream comprising the alkane chosen from the group ofiso-butane, n-butane, cyclopropane and mixtures thereof may be added tothe expanded section via a pipe ring with a plurality of nozzlesdirected to the interior surface of the expanded section.

The liquid-containing stream comprising the alkane chosen from the groupof iso-butane, n-butane, cyclopropane and mixtures thereof may be addedto the expanded section in a circumferential manner, for example byusing at least one nozzle that is tangential to the interior wall of theexpanded section or the liquid containing stream comprising the alkanechosen from the group of iso-butane, n-butane, cyclopropane and mixturesthereof may be added to the expanded section via a pipe ring with aplurality of nozzles directed to the interior surface of the expandedsection.

For example, the liquid-containing stream may be drawn from any pointdownstream of the heat exchanger and before the inlet for the recyclestream and may then be introduced into the expanded section during atleast part of the polymerization process, preferably by bringing theliquid-containing stream is brought into contact with at least part ofthe interior surface of the expanded section.

In a special embodiment, the liquid-containing stream is drawn from anypoint downstream of the heat exchanger and before the inlet for therecycle stream and a gaseous stream is drawn from a point downstream ofthe compressor and upstream of the heat exchanger, after which thegaseous stream and the liquid-containing stream are mixed to form amixture of the gaseous stream and the liquid-containing stream andwherein the mixture of the gaseous stream and the liquid-containingstream is brought into contact with at least part of the interiorsurface of the expanded section.

Preferably, wherein a stream comprising a thermal run away reducingagent (TRRA-containing stream) is introduced into the reactor during atleast part of the polymerization process.

The use of a thermal run away reducing agent in a gas-phase polyolefinpreparation process is known from for example US2009/0118118 A1.

US2009/0118118 A1 relates to a catalyst composition for thepolymerization of propylene comprisigng one or more Ziegler-Nattaprocatalyst compositions comprising one or more esters of aromaticdicarboxylic acid internal electron donors; one or more aluminumcontaining cocatalysts; a selectivity control agent (SCA) comprising atleast one silicon containing compound containing at least one C1-10alkoxy group bonded to a silicon atom, and one or more activity limitingagent (ALA) compounds comprising one or more aliphatic or cycloaliphaticcarboxylic acids; alkyl-, cycloalkyl-, oralkyl(poly)(oxyalkyl)-(poly)ester derivatives thereof; or inertlysubstituted derivatives of the foregoing.

This SCA/ALA mixture is added to the polymerization reactor.

U.S. Pat. No. 5,414,063 discloses a process control scheme fortemporarily slowing or killing the production of polypropylene whileidling the reactor using a fourth generation catalyst system containinga solid magnesium chloride supported titanium tetrachloride catalystcomponent containing an internal strong electron donor in combinationwith a silicon compound, and an organoaluminum compound (cocatalyst) byintroducing an external strong electron donor, for example p-ethoxyethyl benzoate (PEEB) into the reactor.

In the process of the invention, a stream comprising a thermal run awayreducing agent (TRRA-containing stream) may be introduced into thereactor during at least part of the polymerization process. The TRRA maybe introduced into several parts of the reaction system, for example theTRRA may be introduced into the expanded section and/or to the fluidizedbed, for example during at least part of the polymerization process. Forexample one or more TRRA-containing streams may be introduced into theexpanded section during the entire polymerization process. The additionof a TRRA-containing stream may be intermittent or continuous.

TRRA

With thermal run away reducing agent (TRRA) is meant a chemical that iscapable of slowing down or even killing the polymerization reaction byslowing down or even killing the catalyst.

The thermal run away reducing agent (TRRA) may be chosen from the groupconsisting of esters, amines, nitriles, amides and mixtures thereof,preferably from the group of (aromatic) amides and aromatic carboxylicacid esters, more preferably the TRRA is chosen from the group ofaromatic carboxylic acid esters, most preferably, the TRRA is p-ethoxyethyl benzoate (PEEB).

Suitable esters include but are not limited to carboxylic acid esters,for example aliphatic carboxylic acid esters, for example ethylacrylate, methyl methacrylate, di-methyl carbonate, ethylcyclohexanoatem propyl pivalate; and

aromatic carboxylic acid esters, for example ethyl benzoate,methylbenzoate, p-methoxy ethylbenzoate, p-ethoxy methylbenzoate,methylbenzoate, p-ethoxy ethylbenzoate, dimethyloxalate,p-chloroethylbenzoate, p-amino hexylbenzoate, isopropyl naphthanate,n-amyl toluate,

Suitable amines include but are not limited to aminobenzoate, aliphaticamines, for example N,N,N′N′-tetramethyl ethylene diamine;cycloaliphatic amines, for example 1,2,4-trimethyl piperazine,2,3,4,5-tetraethyl piperidine and phthalates, for example dimethylphthalate, diethyl phthalate, di-n-propyl phthalate, diisopropylphthalate, di-n-butyl phthalate, diisobutyl phthalate, di-tert-butylphthalate, diisoamyl phthalate, di-tert-amyl phthalate, dineopentylphthalate, di-2-ethylhexyl phthalate, and di-2-ethyldecyl phthalate.

Suitable nitriles include but are not limited to aromatic and aliphaticnitriles, for example alkane nitriles, for example trimethylacetonitrile.

Suitable amides include but are not limited to aromatic and aliphaticamides, for example n, n-dimethyl benzamide,

Preferably, the TRRA is chosen from the group of esters, for example anaromatic carboxylic acid ester, preferably the TRRA is p-ethoxy ethylbenzoate (PEEB); amides, for example n,n-dimethylbenzamide; or nitriles,for example trimethylacetonitrile.

Preferably, in the process of the invention the thermal run awayreducing agent (TRRA) is chosen from the group consisting of esters,amines, nitriles, amides and mixtures thereof, preferably from the groupof (aromatic) amides and aromatic carboxylic acid esters, morepreferably the TRRA is chosen from the group of aromatic carboxylic acidesters, most preferably, the TRRA is p-ethoxy ethyl benzoate (PEEB).

In the process of the invention, a TRRA-containing stream may be broughtinto contact with at least part of the interior surface of the expandedsection, preferably into contact with the interior wall of the externalsection.

Preferably, in the process of the invention, a TRRA-containing stream isadded to the expanded section in a circumferential manner, for exampleby using at least one nozzle that is tangential to the interior wall ofthe expanded section. Alternatively, a TRRA-containing stream may beadded to the expanded section in a circumferential manner using a duct.

In this way, the expanded section will act as a cyclone for theTRRA-containing stream. (See also FIG. 4, for an example of the path ofthe TRRA-containing stream within the expanded section).

Alternatively, a TRRA-containing stream may be added to the expandedsection via a pipe ring with a plurality of nozzles directed to theinterior surface of the expanded section.

The optimal concentration of TRRA can easily be determined by the personskilled in the art and may depend on the polymerization conditions andcatalyst used and intended polymer. For example, the molar ratio ofTRRA/Si (in the catalyst) may be at least 0.5, for example at least 1,for example at least 1.5, for example at least 2, for example at least2.5, for example at least 3, for example at least 3.5 and/or for exampleat most 6, for example at most, 5.5, for example at most 5, for exampleat most 4.5.

In one embodiment of the invention, a liquid-containing stream is drawnfrom any point downstream of the heat exchanger (b1) and before theinlet for the recycle stream and wherein the alkane chosen from thegroup of iso-butane, n-butane, cyclopropane and mixtures thereof andoptionally a TRRA is/are added to the liquid-containing stream. Anexample of this embodiment is also illustrated by FIG. 1.

Preferably, in the process of the invention, the circulating fluids arecompressed using a compressor and subsequently cooled using a heatexchanger to form the cooled recycle stream.

In such process, preferably a gaseous stream is drawn from a pointdownstream of the compressor and upstream of the heat exchanger.

In a special embodiment of a process wherein the circulating fluids arecompressed using a compressor and subsequently cooled using a heatexchanger to form the cooled recycle stream, the process furthercomprises drawing a liquid-containing stream from any point downstreamof the heat exchanger and before the inlet for the recycle stream anddrawing a gaseous stream from a point downstream of the compressor andupstream of the heat exchanger,

wherein the gaseous stream and the liquid-containing stream are mixed toform a mixture of the gaseous stream and the liquid-containing stream

and wherein the alkane chosen from the group of iso-butane, n-butane,cyclopropane and mixtures thereof and optionally a TRRA is/are added tothe mixture of the gaseous stream and the liquid stream to form a streamcomprising an alkane chosen from the group of iso-butane, n-butane,cyclopropane and mixtures thereof and optionally a TRRA.

The alkane chosen from the group of iso-butane, n-butane, cyclopropaneand mixtures thereof and optionally a TRRA may be added to the gaseousstream and/or to the liquid-containing stream and/or to the mixture ofthe gaseous stream and the liquid-containing stream. Preferably thealkane chosen from the group of iso-butane, n-butane, cyclopropane andmixtures thereof and optionally a TRRA is/are added to the mixture ofthe gaseous stream and the liquid-containing stream.

The advantage of mixing the gaseous and liquid-containing stream is, incase of tangential addition of the stream comprising an alkane chosenfrom the group of iso-butane, n-butane, cyclopropane and mixturesthereof and optionally a TRRA, to ensure the circumferential movement ofthe fluid on the interior wall of the expanded section (which willreduce sheeting in the expanded section).

When a TRRA is added to the fluidized bed, the TRRA is preferably addedto the fluidized bed at a location near where a selectivity controlagent (component which controls the catalyst activity) is added.Examples of selectivity control agents are given in US2009/0118118 A1,hereby incorporated by reference.

Multi-zone Reactor

In a special embodiment, the reactor is a multi-zone reactor asdescribed in patent applications WO2015/078816 and WO2015/078815, herebyincorporated by reference.

In one aspect in the process of the invention, the reactor is amulti-zone reactor suitable for the continuous fluidized bedpolymerization of one or more α-olefin monomers of which at least one isethylene or propylene, which multi-zone reactor is operable in condensedmode, which multi-zone reactor comprises a first zone, a second zone, athird zone, a fourth zone and a distribution plate,

wherein the first zone is separated from the second zone by thedistribution plate, wherein the multi-zone reactor is extended in thevertical direction

wherein the second zone of the multi-zone reactor is located above thefirst zone and wherein the third zone of the multi-zone reactor islocated above the second zone, and wherein the fourth zone of themulti-zone reactor is located above the third zone

wherein the second zone contains an inner wall, wherein at least part ofthe inner wall of the second zone is either in the form of a graduallyincreasing inner diameter or a continuously opening cone, wherein thediameter or the opening increases in the vertical direction towards thetop of the multi-zone reactor

wherein the third zone contains an inner wall, wherein at least part ofthe inner wall of the third zone is either in the form of a graduallyincreasing inner diameter or a continuously opening cone, wherein thediameter or the opening increases in the vertical direction towards thetop of the multi-zone reactor

wherein the largest diameter of the inner wall of the third zone islarger than the largest diameter of the inner wall of the second zone.

In some embodiments, the reactor of the invention may thereby preferablycomprise at least a part of said second zone and/or said third zonecontains an inner wall, wherein at least part of the inner wall has acylindrical shape. The inner wall of the reactor may be the innerenvelope delimiting the reactor.

In the context of the present invention, a gradually increasing diametermay for example mean an increase of the diameter of the inner wall ofthe reactor in the vertical direction towards the top of the reactor.Said increase may be for example stepwise, constant, logarithmic orexponential. One example of such is a continuously opening cone.

In the context of the present invention, a continuously opening cone mayfor example mean a conically shaped part of the inner wall of thereactor comprising a first circular opening and a second circularopening connected via the inner wall of the reactor, in which thederivative of the diameter variation of the wall as measured in thevertical direction towards the top of the reactor may preferably have aconstant and positive value.

In some embodiments of the invention, the zone, preferably for examplethe second zone, in the area directly above the distribution plate iseither in the form of a gradually increasing inner diameter or acontinuously opening cone. In the context hereof, directly above maymean for example that a zone in the form of a gradually increasing innerdiameter or a continuously opening cone, wherein the diameter or theopening increases in the vertical direction towards the top of themulti-zone reactor is located relative to the distribution plate, sothat accumulation of liquids on the surface of the distribution platemay preferably be prevented.

With ‘multi-zone reactor suitable for the continuous polymerization ofone or more α-olefin monomers of which at least one is ethylene orpropylene’ is meant a device capable of containing and controlling thepolymerization of the one of more α-olefin monomers and which device cancomprise a fluidized bed. The multi-zone reactor of the invention ispreferably closed off at the top and the bottom by a hemisphere.

The first zone of the multi-zone reactor is separated from the secondzone by a distribution plate, and is located below the second zone ofthe multi-zone reactor.

In the first zone, a separation and distribution of the gas and liquidmay take place, which is the primary function of the first zone. Thefirst zone may further comprise a flow deflector associated with theentry conduit for providing the bottom recycle stream to prevent theaccumulation of solids and liquids in the first zone. Such flowdeflector is for example described in (the figures of) U.S. Pat. No.4,933,149, hereby incorporated by reference.

The second zone contains an inner wall, wherein at least part of theinner wall of the second zone is either in the form of a graduallyincreasing inner diameter or a continuously opening cone, wherein thediameter or the opening increases in the vertical direction towards thetop of the multi-zone reactor. This leads to a variation of thesuperficial gas velocity at least in a part of the second zone, sincesuperficial gas velocity depends on the circular cross-sectional surfaceinside the reactor. This allows to reduce superficial gas velocity inthe vertical direction towards the top of the multi-zone reactor, sothat the average residence time of polymer particles in the second zonecan be increased as a result.

The continuously opening cone or gradually increasing inner diameter ofthe second zone is preferably located in the lower part of the secondzone, more preferably is located directly above the distribution plate.

The second zone may comprise (part of) the fluidized bed where gas phaseor gas-liquid polymerization may take place. The second zone is suitablefor gas-liquid polymerization (under turbulent fluidization conditions).Turbulent fluidization conditions are described in U.S. Pat. No.6,391,985, hereby incorporated by reference.

In one embodiment of the invention, a gas-liquid polymerization isconducted in the second zone and a gas phase polymerization is conductedin the third zone.

The third zone of the multi-zone reactor is located above the secondzone of the multi-zone reactor. The third zone contains an inner wall,wherein at least part of the inner wall of the third zone is either inthe form of a gradually increasing inner diameter of a continuouslyopening cone, wherein the diameter or the opening increases in thevertical direction towards the top of the multi-zone reactor. This leadsto a variation of the superficial gas velocity at least in a part of thethird zone, since superficial gas velocity depends on the circularcross-sectional surface inside the reactor. This allows to reducesuperficial gas velocity in the vertical direction towards the top ofthe multi-zone reactor, so that the average residence time of polymerparticles in the second zone can be increased as a result.

By using the multi-zone reactor in the process of the invention, in thesecond zone (2) a gas-liquid polymerization may take place and in thethird zone, a gas-phase polymerization may then occur. Therefore, theinvention may provide a two-stage polymerization.

The top zone or fourth zone is the expanded section.

The third zone may comprise part of the fluidized bed. The third zone issuitable for gas-phase polymerization.

The third zone and the second zone can be distinguished when themulti-zone reactor is operated; however there is no sharp boundarybetween the second and third zone. Typically, when operating themulti-zone reactor, the second zone will comprise more liquid than thethird zone and in the third zone, a gas-phase polymerization will takeplace.

In such multi-zone reactor, during the course of polymerization, freshpolymer particles are produced by catalytic polymerization of α-olefinmonomers. Such polymer particles are projected upwards in the directionof the fourth zone through the circulating gas. Most of these particlesdo preferably not reach the fourth zone or return to the second or thirdzone by gravity as the superficial velocity decreases in the fourthzone. The fourth zone may be connected to the third zone or optionalfurther zone(s).

The multi-zone reactor of the invention may comprise further zones, suchas for example one, two or even optionally three further zones, that canfor example be a fifth zone and optionally a sixth zone and optionallyeven a seventh zone. These zones may provide a further possibility forpolymerization, wherein each further zone may be operated at differentreaction conditions. These further zones can be located preferablybetween the third zone and the top zone.

With inner diameter is meant the diameter in a given horizontal planeperpendicular to the center line of the multi-zone reactor as measuredfrom the inside of the inner wall of the multi-zone reactor.

For example, the maximum inner diameter of the fourth zone is at least1, for example at least 3, for example at least 5% and/or for example atmost 300%, for example at most 200%, for example at most 150%, forexample at most 80%, for example at most 70%, for example at most 60%,for example at most 50%, for example at most 40%, for example at most30%, for example at most 25%, for example at most 20%, for example atmost 15% larger than the maximum inner diameter of the third zone. Forexample, the maximum inner diameter of the fourth zone is from 5 to 30%larger than the maximum inner diameter of the third zone.

For example, the maximum inner diameter of the third zone is at least 1,for example at least 3, for example at least 5% and/or for example atmost 300%, for example at most 200%, for example at most 150%, forexample at most 80%, for example at most 70%, for example at most 60%,for example at most 50%, for example at most 40%, for example at most30%, for example at most 25%, for example at most 20%, for example atmost 15% larger than the maximum inner diameter of the second zone. Forexample, the maximum inner diameter of the third zone is from 15 to 30%larger than the maximum inner diameter of the second zone.

For example, the maximum inner diameter of the second zone is at least1, for example at least 3, for example at least 5% and/or for example atmost 300%, for example at most 200%, for example at most 150%, forexample at most 80%, for example at most 70%, for example at most 60%,for example at most 50%, for example at most 40%, for example at most30%, for example at most 25%, for example at most 20%, for example atmost 15% larger than the maximum inner diameter of the first zone. Forexample, the maximum inner diameter of the second zone is from 15 to 30%larger than the maximum inner diameter of the first zone.

In one embodiment, the invention relates to the reactor of theinvention, wherein at least the bottom part of the third zone comprisesan inner wall in the form of a gradually increasing inner diameter or acontinuously opening cone, wherein the diameter or the opening increasesin the vertical direction towards the top of the multi-zone reactor. Inthis embodiment, the bottom part of the second zone and/or of the bottompart of the fourth zone may also comprise an inner wall in the form of agradually increasing inner diameter or a continuously opening cone,wherein the diameter or the opening increases in the vertical directiontowards the top of the multi-zone reactor.

In one embodiment, as illustrated in FIG. 3 representing the embodimentof addition of a stream comprising an akane chosen from the group ofiso-butane, n-butane, cyclopropane and mixtures thereof, of FIG. 2, thezone (2) in the area directly above the distribution plate is either inthe form of a gradually increasing inner diameter or a continuouslyopening cone (2A), wherein the diameter or the opening increases in thevertical direction towards the top of the multi-zone reactor and whereinthe top part of the second zone has an inner wall having a cylindricalshape (2B) and wherein the top part of the second zone is connected to abottom part of the third zone (3A), wherein the bottom part of the thirdzone is either in the form of a gradually increasing inner diameter or acontinuously opening cone, wherein the diameter or the opening increasesin the vertical direction towards the top of the multi-zone reactor andwherein the top part of the third zone has an inner wall having acylindrical shape (3B) and wherein the top part of the third zone isconnected to the top zone, for example to the fourth zone.

Therefore, preferably, zone (2) in the area directly above thedistribution plate is either in the form of a gradually increasing innerdiameter or a continuously opening cone (2A), wherein the diameter orthe opening increases in the vertical direction towards the top of themulti-zone reactor and wherein the top part of the second zone has aninner wall having a cylindrical shape (2B) and wherein the top part ofthe second zone is connected to a bottom part of the third zone (3A),wherein the bottom part of the third zone is either in the form of agradually increasing inner diameter or a continuously opening cone,wherein the diameter or the opening increases in the vertical directiontowards the top of the multi-zone reactor and wherein the top part ofthe third zone has an inner wall having a cylindrical shape (3B) andwherein the top part of the third zone is connected to the top zone, forexample to the fourth zone.

Preferably, the cylindrical shape is the shape of a right circularcylinder.

Preferably, the angle (a) of the inner wall of the part of the secondzone having the gradually increasing inner diameter or having thecontinuously opening cone, relative to the centre line (9) of themulti-zone reactor (8) is from 0.1 to 80 degrees, preferably from 1 to60 degrees, more preferably from 1-45 degrees, most preferably around 27degrees.

For example, said angle (a) is at least 5, for example at least 7, forexample at least 10 degrees, for example at least 20 degrees and/or forexample at most 60, for example at most 50, for example at most 40, forexample at most 35 degrees, for example at most 30 degrees. For example,the angle (a) is in the range from 10 to 40 degrees.

Preferably, the invention uses a multi-zone reactor as described herein,wherein the angle (a) of the inner wall of the part of the third zonehaving the gradually increasing inner diameter or having thecontinuously opening cone, relative to the centre line (9) of themulti-zone reactor (8) is from 0.1 to 80 degrees, preferably from 1 to60 degrees, more preferably from 1-45 degrees, most preferably around 27degrees, for example from 1 to 40 degrees.

For example, said angle (a) is at least 5, for example at least 7, forexample at least 10 degrees, for example at least 20 degrees and/or forexample at most 60, for example at most 50, for example at most 40, forexample at most 35 degrees, for example at most 30 degrees. For example,the angle (a) is in the range from 10 to 40 degrees.

It should be appreciated by the skilled person that due to the fact thatthe volume in the multi-zone reactor of the invention expands from thefirst zone to the second zone and from the second zone to the third zoneand from the third zone to the fourth zone when operating the multi-zonereactor, the superficial gas velocities in these zones will decreasefrom the first to the second and from the second to the third zone andfrom the third zone to the fourth zone. For example, the superficial gasvelocities in the zones of the multi-zone reactor of the invention whenused to produce polypropylene, for example homo polypropylene, may be inthe range of from 0.3 to 3 m/s, for example in the range from 0.4 to 2m/s.

Some special embodiments of systems suitable for the process of theinvention are schematically represented in FIGS. 1-4 without howeverbeing limited thereto. The system of FIG. 1 (FIG. 1) is only one ofnumerous possible schematic arrangements. Thus, for example, thesequence of the equipment items in the circulated gas line, particularlyof the cooler and compressor can also be reversed or further equipmentitems can be integrated into the line. Further elements such as systemsfor metering-in the catalyst are not shown in FIG. 1, such elements areknown to those skilled in the art and can be integrated into the reactorin a known manner.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 illustrates a system suitable for the continuous preparation of apolyolefin according to the process of the invention.

FIG. 2 illustrates a special embodiment of the system of FIG. 1.

FIG. 3 illustrates a special embodiment of the system of FIG. 1, whereinthe reactor is a multi-zone reactor.

FIG. 4 illustrates a special embodiment of the addition of an alkanechosen from the group of iso-butane, n-butane, cyclopropane and mixturesthereof to the expanded section of the reactor.

FIG. 5 is a TREF profile of the polymer produced in example 2.

FIG. 6 is a TREF profile of the polymer produced in example 3.

FIG. 1 illustrates a system suitable for the continuous preparation of apolyolefin in a reactor from one or more α-olefin monomers of which atleast one is ethylene or propylene, comprising a reactor (8), acompressor (400), a heat exchanger (5), an expanded section (4) (whichexpanded section is located at or near the top of the reactor)

wherein the reactor comprises

-   -   a distribution plate (6) (which distribution plate is located at        the lower part of the reactor) and    -   an inlet for a cooled recycle stream (10) located under the        distribution plate    -   an inlet for providing the catalyst (20)    -   an outlet for providing the polyolefin (30)    -   an outlet for a recycle stream (40),

wherein the outlet for the recycle stream (40) of the expanded sectionis connected to an inlet of the compressor (400) via a first connectionmeans (AA), for instance pipes

wherein the compressor (400) comprises an outlet for compressed fluids(50),

wherein the outlet of the compressor (400) is connected to an inlet forcompressed fluids of the heat exchanger (5) via a second connectionmeans (BB)

wherein optionally the second connection means (BB), for instance pipes,comprises an inlet for receiving a feed (70),

wherein the heat exchanger (5) comprises an outlet for providing thecooled recycle stream (10) which outlet of the heat exchanger (5) isconnected to the inlet of the reactor (8) for receiving the cooledrecycle stream (10),

wherein the first connection means (AA) may comprise an inlet forreceiving a feed (60) and further comprising means for adding an alkanechosen from the group of iso-butane, n-butane, cyclopropane and mixturesthereof and optionally a TRRA (a) to the reactor such that the molarcomposition of the components in the recycle stream is chosen such thatthe dew temperature of the recycle stream at the reactor pressure is atleast 0.1° C. below the temperature of the reactor

In this special embodiment of the system of the invention, an alkanechosen from the group of iso-butane, n-butane, cyclopropane and mixturesthereof and optionally a TRRA (a) is added to a liquid-containing streamdrawn from any point downstream of the heat exchanger (b1) and/or analkane chosen from the group of iso-butane, n-butane, cyclopropane andmixtures thereof is added to the reactor by addition of this stream to agaseous stream that is drawn from a point downstream of the compressor(400) and upstream of the heat exchanger (5) (b2) and wherein the systemfurther may comprise one or more inlets for receiving an alkane chosenfrom the group of iso-butane, n-butane, cyclopropane and mixturesthereof and optionally a TRRA (a).

FIG. 2 illustrates a preferred embodiment of the system of FIG. 1 andshows that the gaseous stream (b2) and the liquid-containing stream (b1)are connected to a mixing unit (3) which mixing unit is connected to themeans for receiving the an alkane chosen from the group of iso-butane,n-butane, cyclopropane and mixtures thereof (b5) and wherein the systemfurther comprises one or more inlets for receiving an alkane chosen fromthe group of iso-butane, n-butane, cyclopropane and mixtures thereof andoptionally a TRRA (a).

FIG. 3 illustrates a a special embodiment of the system of FIG. 1,wherein the reactor is a multi-zone reactor,

which multi-zone reactor is suitable for the continuous fluidized bedpolymerization of one or more α-olefin monomers of which at least one isethylene or propylene, which multi-zone reactor is operable in condensedmode, which multi-zone reactor comprises a first zone, a second zone, athird zone, a fourth zone and a distribution plate,

wherein the first zone is separated from the second zone by thedistribution plate, wherein the multi-zone reactor is extended in thevertical direction

wherein the second zone of the multi-zone reactor is located above thefirst zone and wherein the third zone of the multi-zone reactor islocated above the second zone, and wherein the fourth zone of themulti-zone reactor is located above the third zone

wherein the second zone contains an inner wall, wherein at least part ofthe inner wall of the second zone is either in the form of a graduallyincreasing inner diameter or a continuously opening cone, wherein thediameter or the opening increases in the vertical direction towards thetop of the multi-zone reactor

wherein the third zone contains an inner wall, wherein at least part ofthe inner wall of the third zone is either in the form of a graduallyincreasing inner diameter or a continuously opening cone, wherein thediameter or the opening increases in the vertical direction towards thetop of the multi-zone reactor

wherein the largest diameter of the inner wall of the third zone islarger than the largest diameter of the inner wall of the second zone,

wherein the zone (2) in the area directly above the distribution plateis either in the form of a gradually increasing inner diameter or acontinuously opening cone (2A), wherein the diameter or the openingincreases in the vertical direction towards the top of the multi-zonereactor and wherein the top part of the second zone has an inner wallhaving a cylindrical shape (2B) and wherein the top part of the secondzone is connected to a bottom part of the third zone (3A), wherein thebottom part of the third zone is either in the form of a graduallyincreasing inner diameter or a continuously opening cone, wherein thediameter or the opening increases in the vertical direction towards thetop of the multi-zone reactor and wherein the top part of the third zonehas an inner wall having a cylindrical shape (3B) and wherein the toppart of the third zone is connected to the top zone, for example to thefourth zone.

In FIG. 1 and, FIG. 2 and FIG. 4, the fluidized bed (80) is indicatedwith dots.

FIG. 4 illustrates the flow of the liquid containing stream (b2) towhich an alkane chosen from the group of iso-butane, n-butane,cyclopropane or mixtures thereof has been added in the expanded section,when such stream is added to the expanded section in a circumferentialmanner. In this figure, it is shown how the use of a nozzle that istangential to the interior wall of the expanded section forces the(centrifugal/circumferential) movement of the stream along the interiorwalls.

In another aspect, the invention relates to a reaction system for thepreparation of polypropylene from propylene and optionally at least oneother α-olefin monomer,

wherein the reaction system comprises a reactor (8),

wherein the reactor comprises a fluidized bed (80), an expanded section(4) located at or near the top of the reactor, a distribution plate (6)located at the lower part of the reactor and an inlet for a recyclestream (10) located under the distribution plate (6)

wherein the system is arranged such that

-   -   a polymerization catalyst (20) is fed to the fluidized bed (80)        in the area above the distribution plate (6)    -   the propylene and the optional at least one other α-olefin        monomer are fed to the reactor (8)    -   the polypropylene (30) is withdrawn from the reactor    -   fluids (40) are circulated from the top of the reactor to the        bottom of the reactor, wherein the circulating fluids are        compressed using a compressor (400) and cooled using a heat        exchanger (5), resulting in a cooled recycle stream comprising        liquid (10), and wherein the cooled recycle stream is introduced        into the reactor using the inlet for the recycle stream

wherein an alkane chosen from the group of iso-butane, n-butane,cyclopropane and mixtures and optionally a TRRA (a) thereof is added tothe reactor and

wherein the molar composition of the components in the recycle stream ischosen such that the dew temperature of the recycle stream at thereactor pressure is at least 0.1° C. below the temperature of thereactor.

Preferably, the reaction system is further arranged that a streamcomprising a thermal run away reducing agent (TRRA-containing stream) isintroduced into the reactor during at least part of the polymerizationprocess.

In the process and system of the invention, the one or more α-olefinmonomers and other fluids, such as hydrogen, an inert gas or liquid, forexample a condensable inert component, may be added to the recyclestream (40) to make up for reacted fluids before cooling the fluids toform the cooled recycle stream.

The feed (60) comprises a chain transfer agent, for example hydrogen andmay further comprise gaseous α-olefin monomers and insert gaseouscomponents, for example nitrogen. A chain transfer agent, such ashydrogen may for instance be used to adjust the molecular weight of thepolyolefin (30) produced.

The feed (70) comprises condensable inert components, for example acondensable inert component selected from the group of alkanes having 4to 20 carbon atoms, preferably 4 to 8 carbon atoms, and mixturesthereof, for example propane, n-butane, isobutene, n-pentane,isopentane, neopentane, n-hexane, isohexane or other saturatedhydrocarbons having 6 C-atoms, n-heptane, n-octane and other saturatedhydrocarbons having 7 or 8 C-atoms and any mixtures thereof; and mayfurther comprise condensable α-olefin monomers, α-olefin comonomersand/or mixtures thereof.

The condensable inert component is preferably selected from the group ofisopentane, n-hexane, n-butane, i-butane and mixtures thereof. Becauseof their more attractive pricing, preferably isopentane and/or n-hexaneare/is used as condensable inert component(s) in the feed (70)

When copolymers are produced, the process of the invention furthercomprises supplying a comonomer using feed (60) or (70) in case of anon-condensable comonomer and using feed (70) in case of a condensablecomonomer.

It is apparent to the skilled person that the process of the presentinvention may also be applied when using multiple reactor. For example,for the purpose of the present invention, it is to be understood that ifmultiple reactors are employed, the alkane chosen from the group ofiso-butane, n-butane, cyclopropane and mixtures thereof may be added tothe expanded section of any one of the reactors—and in case of tworeactors, either reactor—and that the alkane need not be added to allreactors of the multiple reactor train.

With ‘condensed mode’ is meant that a liquid containing stream is usedto cool the multi-zone reactor (8).

Hydrogen may for instance be used as a chain transfer agent to adjustthe molecular weight of the polyolefin (30) produced.

It is apparent to the skilled person that recycle streams may be presentin the reaction system of the invention, for example there may be arecycle stream that is vented back from a polymer discharge system tothe fluid bed reactor aiming at an efficient discharge of the productwhile at the same time recycling a large portion of unreacted gassesback to the reactor.

The continuous polymerization of one or more α-olefin monomers willproduce polyolefins in the form of particles, herein also referred to as‘polyolefin’ (30). Examples of polyolefins which may thus produced,include a wide variety of polymers, for example polyethylene, forexample linear low density polyethylene (LLDPE), which may for examplebe prepared from ethylene and but-1-ene, 4-methylpent-1-ene orhex-1-ene, high density polyethylene (HDPE), which may for example beprepared from ethylene or from ethylene with a small portion of anα-olefin monomer having from 4 to 8 carbon atoms, for example but-1-ene,pent-1-ene, hex-1-ene or 4-methylpent-1-ene. Other examples include butare not limited to plastomers, elastomers, medium density polyethylene,polypropylene, such as propylene homopolymers and propylene copolymers,including random copolymers of propylene and ethylene or randomcopolymers of propylene and α-olefin monomer having from 4 to 8 carbonatoms, for example but-1-ene, pent-1-ene, hex-1-ene or4-methylpent-1-ene and optionally ethylene, for example a randomcopolymer of propylene and at most 5 wt % ethylene, and block ormulti-block copolymers of propylene and optionally a comonomer (such asethylene and/or α-olefin monomer having from 4 to 8 carbon atoms) andethylene propylene rubber (EPR).

Preferably, in the process of the invention, the polyolefin produced isa propylene, for example a propylene homopolymer, a propylene copolymeror a heterophasic copolymer, that is a composition comprising apropylene homopolymer or a propylene copolymer and an ethylene propylenerubber.

In another aspect therefore, the invention relates to a polypropyleneobtained or obtainable by the process of the invention.

It is further noted that the term ‘comprising’ does not exclude thepresence of other elements. However, it is also to be understood that adescription on a product comprising certain components also discloses aproduct consisting of these components. Similarly, it is also to beunderstood that a description on a process comprising certain steps alsodiscloses a process consisting of these steps.

The invention will now be elucidated by way of the following examples,without however being limited thereto.

EXAMPLE 1

A computer-based mathematical model capable of generating mass and heatbalances along a fluidized bed reactor was used to run simulation incondensed mode operation to show the advantage of adding an alkanehaving 3 to 6 carbon atoms in the process of the invention. Firstly, themodel was run using actual data from commercial polypropylene productionto validate the model. The results are shown in Table 1.

TABLE 1 Commercial data validated versus the computer-based mathematicalmodel Example Number b (Model Reactor Conditions a (Commercial) Results)Internal Reactor diameter (m) 4.8 4.8 Recycle Gas Superficial Velocity(m/s) 0.3081 0.3081 Recycle Gas Composition (mole fraction): Propylene0.83893 0.83893 Propane 0.09398 0.09398 Hydrogen 0.00267 0.00267Nitrogen 0.06443 0.06443 n-Butane — — Iso-Butane — — Recycle Gas Density(kg/m³.) 67.25 72.31 Reactor Temperature (° C.) 68.08 68.08 ReactorInlet Temperature (° C.) 65.70 65.70 Reactor Pressure (kPag) 3214.853214.85 Reactor inlet Pressure (kPag) 3334.53 3334.53 Inlet Dew PointTemperature (° C.) 67.19 68.56 Condensed Liquid in Recycle 39.78 36.94Stream (% weight) Production Rate (ton/h) 50.6 51.2

As can be seen from Table 1, the actual data and the data from the modelare very well comparable.

Subsequently, the model was ran representing commercial data when analkane having 3 to 6 carbon atoms was not added (propane present in therecycle stream is produced by hydrogenation of propylene in the reactorin case hydrogen is present and/or comes from the feed-stream comprisingpropylene) versus when different amounts of n-butane, iso-butane orcyclopropane were added to the recycle stream. The n-butane, i-butane orcyclopropane were added using feed (70)

The results are shown in Table 2 below.

As can be seen from Table 2, the feeding of an alkane having 3-6 carbonatoms increases the production rate from 51.2 tons/hour to for 2.0 mol %n-butane to 52.9 tons/hour, for 2.3 mol % n-butane to 54.9 tons/hour;for 2.65 mol % i-butane to 69.4 tons/hour, for 2.75 mole % i-butane to71 tons/hour and for 3.89 mol % cyclopropane to 73.1 tons/hour.

Therefore, it has been shown that in the process of the invention, it isadvantageous to add an wherein an alkane chosen from the group ofiso-butane, n-butane, cyclopropane and mixtures thereof is added to thereactor and wherein the molar composition of the components in therecycle stream is chosen such that the dew temperature of the recyclestream at the reactor pressure is at least 0.1° C. below the temperatureof the reactor.

The effect of the TRRA is demonstrated by the examples 2 and 3 below;

These examples show that TRRAs, for example esters, amines, nitriles,amides and mixtures thereof, in particular paraethoxyethylbenzoate(PEEB), trimethylacetonitrile (TA) and n,n-dimethyl benzamide (DB) arecapable of reducing the catalyst activity (and hence controlling thetemperature in the expanded section) without however affecting theproperties of the produced polyolefin.

TABLE 2 Reactor Conditions b I II III IV V n-butane (mol %) — 2.0 2.3i-butane (mol %) — — — 2.65 2.75 cyclopropane 3.89 (mol %) InternalReactor 4.8 4.8 4.8 4.8 4.8 4.8 diameter (m) Recycle Gas 0.3081 0.30810.3081 0.3081 0.3081 0.3081 Superficial Velocity (m/s) Recycle GasComposition (mole fraction): Propylene 0.83893 0.80893 0.80593 0.805930.80593 0.80000 Propane 0.09398 0.09398 0.09398 0.09398 0.09398 0.09398Hydrogen 0.00267 0.00267 0.00267 0.00267 0.00267 0.00267 Nitrogen0.06443 0.07443 0.07443 0.07093 0.06993 0.06443 n-Butane — 0.020000.02300 — — — Iso-Butane — — — 0.0265 0.0275 cyclopropane — — — — —0.03893 Recycle Gas 72.31 72.17 72.41 73.12 73.38 73.17 Density (kg/m³.)Reactor 68.08 68.08 68.08 68.08 68.08 68.08 Temperature (° C.) ReactorInlet 0.03893 0.03893 0.03893 0.03893 0.03893 65.70 Temperature (° C.)Reactor Pressure 3214.85 3214.85 3214.85 3214.85 3214.85 3214.85 (kPag)Reactor inlet 3334.53 3334.53 3334.53 3334.53 3334.53 3334.53 Pressure(kPag) Inlet Dew Point 68.56 69.44 69.73 71.03 71.03 70.69 Temperature(° C.) Condensed Liquid 36.94 38.6 40.18 51.10 52.25 53.40 in RecycleStream (% weight) Production Rate 51.2 52.9 54.9 69.4 71.0 73.1 (ton/h)

EXAMPLE 2 TRRA is PEEB

The polymerization catalyst was prepared as follows:

EXAMPLE 2 Preparation of a Procatalyst on an Activated Butyl-GrignardSupport

Preparation of Grignard Reagent (Step o))—Phase A

This step o) constitutes the first part of phase A of the process forpreparation of the procatalyst.

A stirred flask, fitted with a reflux condenser and a funnel, was filledwith magnesium powder (24.3 g). The flask was brought under nitrogen.The magnesium was heated at 80° C. for 1 hour, after which dibutyl ether(150 ml), iodine (0.03 g) and n-chlorobutane (4 ml) were successivelyadded. After the colour of the iodine had disappeared, the temperaturewas raised to 80° C. and a mixture of n-chlorobutane (110 ml) anddibutyl ether (750 ml) was slowly added for 2.5 hours. The reactionmixture was stirred for another 3 hours at 80° C. Then the stirring andheating were stopped and the small amount of solid material was allowedto settle for 24 hours. By decanting the colourless solution above theprecipitate, a solution of butylmagnesiumchloride with a concentrationof 1.0 mol Mg/I was obtained.

Preparation of Solid Magnesium Compound (Step i))—Phase A

This step i) constitutes the second part of phase A of the process forpreparation of the procatalyst.

This step is carried out as described in Example XX of EP 1 222 214 B1,except that the dosing temperature of the reactor is 35° C., the dosingtime is 360 min and the propeller stirrer w is as used. An amount of 250ml of dibutyl ether is introduced to a 1 liter reactor. The reactor isfitted by propeller stirrer and two baffles. The reactor is thermostatedat 35° C.

The solution of reaction product of step A (360 ml, 0.468 mol Mg) and180 ml of a solution of tetraethoxysilane (TES) in dibutyl ether (DBE),(55 ml of TES and 125 ml of DBE), are cooled to 10° C., and then aredosed simultaneously to a mixing device of 0.45 ml volume supplied witha stirrer and jacket. Dosing time is 360 min. Thereafter the premixedreaction product A and the TES-solution are introduced to a reactor. Themixing device (minimixer) is cooled to 10° C. by means of cold watercirculating in the minimixer's jacket. The stirring speed in theminimixer is 1000 rpm. The stirring speed in reactor is 350 rpm at thebeginning of dosing and is gradually increased up to 600 rpm at the endof dosing stage.

On the dosing completion the reaction mixture is heated up to 60° C. andkept at this temperature for 1 hour. Then the stirring is stopped andthe solid substance is allowed to settle. The supernatant is removed bydecanting. The solid substance is washed three times using 500 ml ofheptane. As a result, a pale yellow solid substance, reaction product B(the solid first intermediate reaction product; the support), isobtained, suspended in 200 ml of heptane. The average particle size ofsupport is 22 μm and span value (d₉₀-d₁₀)/d₅₀=0.5.

Activation of First Intermediate Reaction Product (Step ii))—Phase B

This step ii) constitutes phase B of the process for preparation of theprocatalyst as discussed above.

Support activation was carried out as described in Example IV ofWO/2007/134851 to obtain the second intermediate reaction product.

In inert nitrogen atmosphere at 20° C. a 250 ml glass flask equippedwith a mechanical agitator is filled with slurry of 5 g of reactionproduct B dispersed in 60 ml of heptane. Subsequently a solution of 0.22ml ethanol (EtOH/Mg=0.1) in 20 ml heptane is dosed under stirring during1 hour. After keeping the reaction mixture at 20° C. for 30 minutes, asolution of 0.79 ml titanium tetraethoxide (TET/Mg=0.1) in 20 ml ofheptane was added for 1 hour.

The slurry was slowly allowed to warm up to 30° C. for 90 min and keptat that temperature for another 2 hours. Finally the supernatant liquidis decanted from the solid reaction product (the second intermediatereaction product; activated support) which was washed once with 90 ml ofheptane at 30° C.

The activated support, according to chemical analysis, comprises amagnesium content of 17.3 wt. %, a titanium content of 2.85 wt. %, and achloride content of 27.1 wt. % corresponding to a molar ratio of Cl/Mgof 1.07 and Ti/Mg of 0.084.

C. Preparation of the Procatalyst

A reactor was brought under nitrogen and 125 ml of titaniumtetrachloride was added to it. The reactor was heated to 90° C. and asuspension, containing about 5.5 g of the support obtained in step C in15 ml of heptane, was added to it under stirring. The reaction mixturewas kept at 90° C. for 10 min. Then ethyl benzoate was added (EB/Mg=0.15molar ratio). The reaction mixture was kept for 60 min. Then thestirring was stopped and the solid substance was allowed to settle. Thesupernatant was removed by decanting, after which the solid product waswashed with chlorobenzene (125 ml) at 90° C. for 20 min. The washingsolution was removed by decanting, after which a mixture of titaniumtetrachloride (62.5 ml) and chlorobenzene (62.5 ml) was added. Thereaction mixture was kept at 90° C. for 30 min. After which the stirringwas stopped and the solid substance was allowed to settle. Thesupernatant was removed by decanting, after which a mixture of titaniumtetrachloride (62.5 ml) and chlorobenzene (62.5 ml) was added. Thendi-n-butyl phthalate (DBP) (DBP/Mg=0.15 molar ratio) in 3 ml ofchlorobenzene was added to reactor and the temperature of reactionmixture was increased to 115° C. The reaction mixture was kept at 115°C. for 30 min. After which the stirring was stopped and the solidsubstance was allowed to settle. The supernatant was removed bydecanting, after which a mixture of titanium tetrachloride (62.5 ml) andchlorobenzene (62.5 ml) was added. The reaction mixture was kept at 115°C. for 30 min, after which the solid substance was allowed to settle.The supernatant was removed by decanting and the solid was washed fivetimes using 150 ml of heptane at 600° C., after which the procatalystIll, suspended in heptane, was obtained.

The Polymerization was Conducted as Follows:

Propylene polymerization experiments (Table 1) were performed usingprocatalysts I, II and Ill described above. Triethylaluminium (TEAL) wasused as co-catalyst, and cyclohexylmethyldimethoxysilane (C-donor) orn-propyltrimethoxysilane (N-donor) was used as external donor (Si).Experiments were performed at different H2/C3 molar ratios.

The polymerization of propylene was carried out in a stainless steel gasphase reactor with a volume of 1800 mL. Under a nitrogen atmosphere, theco-catalyst (TEAL) and procatalyst synthesized according to theprocedure described above and the external electron donor were dosed tothe reactor as heptane solutions or slurries. 10-15 mg (2% wt Ti) ofprocatalyst were employed. The molar ratio of co-catalyst TEAL totitanium (from the procatalyst) was set to 130, and the Si/Ti ratio wasset to 8. During this dosing, the reactor temperature was maintainedbelow 30° C. Subsequently, the reactor was pressurized using a set ratioof propylene and hydrogen, and the temperature and pressure were raisedto its setpoint (67 or 82° C. and 20 barg). After the pressure setpointhas been reached, the polymerization was continued for 60 minutes.During the polymerization reaction the gas cap composition of propyleneand hydrogen was controlled using mass flow meters and online-GCcontrol. After reaching the polymerization time the reactor wasdepressurized and cooled to ambient conditions. The propylene polymer soobtained was removed from the reactor and stored in aluminium bags.

The polymerization conditions are summarized in Table 3 below. In theseexperiments paraethoxyethylbenzoate (PEEB) was used as TRRA. Thetemperature of 67° C. reflects the temperature of a polymerizationwithin the fluidized bed; the temperature of 82° C. reflects thetemperature within the expanded section.

TABLE 3 Experimental polymerization conditions. Parameter At T = 67° C.At T = 82° C. PC3 (barg) 26 28 H2 (mol %) 0.37 0.37 TEAL/Ti (mol/mol)130 130 Si/Ti (mol/mol) 8 8 PC3 pressure of propylene H2 hydrogen TEALtri-ethylaluminium

The effect of different TRRA/Si ratios on the yield is shown in Table 4below:

TABLE 4 yield versus TRRA/Si molar ratio at 67 and 82° C. Molar ratioYield at 67° C. Yield at 82° C. TRRA/Si (Kg-PP/g-Cat) (Kg-PP/g-Cat)TRRA/Si = 0 24.2 20.0 TRRA/Si = 0.68 21.6 16.8 TRRA/Si = 0.98 — 17.2TRRA/Si = 1.12 — 15.4 TRRA/Si = 1.25 — 12.8 TRRA/Si = 2 — 12.5 TRRA/Si =4 17.9 5.9 TRRA/Si = 4.5 19.9 7.6

The polypropylene produced with TRRA/Si ratio 0; 4 and 4.5 andpolymerization temperatures of 67° C. or 82° C. was characterized interms of its molecular weight distribution (MWD) and crystallinity.

Crystallinity was determined using analytical temperature rising elutionfractionation (ATREF) analysis was conducted according to the methoddescribed in U.S. Pat. No. 4,798,081 and Wilde, L.; Ryle, T. R.;Knobeloch, D. C; Peat, L R.; Determination of Branching Distributions inPolyethylene and Ethylene Copolymers, J. Polym. ScL, 20, 441-455 (1982),which are incorporated by reference herein in their entirety. Thecomposition to be analyzed was dissolved in 1,2-dichlorobenzene assolvent of analytical quality filtrated via 0.2 μm filter and allowed tocrystallize in a column containing an inert support (Column filled with150 μm stainless steel beans (volume 2500 μL) by slowly reducing thetemperature to 20° C. at a cooling rate of 0.1° C./min. 1 g/L Irgafosand BHT were used as stabilizers The column was equipped with aninfrared detector. An ATREF chromatogram curve was then generated byeluting the crystallized polymer sample from the column by slowlyincreasing the temperature of the eluting solvent (1,2-dichlorobenzene)from 20 to 130° C. ata rate of 1° C./min.

The instrument used was Polymer Char Crystaf-TREF 300.

-   Stabilizers: 1 g/L Topanol+1 g/L Irgafos 168-   Sample: approx. 70 mg in 20 mL-   Sample volume: 0.3 mL-   Pump flow: 0.50 mL/min

The software from the Polymer Char Crystaf-TREF-300 was used to generatethe spectra.

molecular weight distribution was determined using IAV MolecularCharacterization method. The chromatography equipment used is PolymerLaboratories PL-GPC220 with Viscotek 220R viscometer and Refractiveindex detector. The column set consists of three columns of PolymerLaboratories 13 μm PLgel Olexis, 300×7.5 mm. Standard linearpolyethylene was used for calibration and reference.

The results of the aTREF analysis are represented in FIG. 5.

-   Profile A in FIG. 5 shows the aTREF temperature profile of a    polypropylene produced without a TRRA at T=67° C.-   Profile B in FIG. 5 also shows the aTREF temperature profile of a    polypropylene produced without a TRRA at T=82° C.-   Profile C in FIG. 5 shows the aTREF temperature profile of a    polypropylene produced with TRRA/Si ratio of 4 at T=67° C.

As can be seen in FIG. 5 the addition of a TRRA does not significantlyaffect the crystallinity of the polymer produced at normal operatingtemperature (67° C.).

The crystallinity of these polypropylenes shows an almost similar peaktemperature with a slight shift of the polymer sample produced at 82° C.towards lower crystallinity.

The results of the molecular weight distribution is given in Table 5below:

TABLE 5 Summary of the molecular weight averages and their correspondingviscosities. Polymerization TRRA/Si temperature Mn Mw Mz [η] Run(mol/mol) (° C.) (×10³ g/mol) (×10³ g/mol) (×10³ g/mol) Mw/Mn Mz/Mw(dL/g) 1 0 67 69 460 1500 6.7 3.2 2.10 2 0 82 68 340 1000 5.1 2.9 1.73 34 67 89 550 1700 6.2 3.1 2.42 4 4 82 82 460 1600 5.7 3.5 2.12 5 4.5 8290 470 1300 5.2 2.8 2.17 6 4.5 67 95 560 1700 6.0 3.1 2.46

The results of Table 5 show that a molecular weight distribution (Mw/Mn,MWD) closer to the desired MWD (run #1, MWD=6.7) is obtained when usinga TRRA/Si in a preferred ratio of 4 mol/mol (runs #3 and 4) as comparedto not using a TRRA at 82° C. (run #2). Moreover, since the activity ofthe catalyst is more reduced by the TRRA at a higher temperature, thecontribution of the MWD of the polymer produced at 82° C. is less,thereby making an overall product that has a MWD closer to the desiredMWD of 6.7 (combination of a bit of 5.7 of run #4 and 6.2 of run #2)than in the situation where a TRRA is not used; in the latter case thecontribution of the MWD of 5.1 at 82° C. is more pronounced.

EXAMPLE 3

-   With C-donor is meant: cyclohexylmethyldimethoxysilane.-   With N-donor is meant: n-propyltrimethoxysilane

The polymerization catalyst was prepared as described in Example-1.Similar procedures and equipment used in Example-1 were used in thisexample for the molecular weight distribution (MWD) using GPC andcrystallinity using ATREF.

The effect of different TRRA/Si ratios for both TA and DB on the yieldare shown in Table 6 below:

TABLE 6 Yield versus TRRA/Si molar ratio for TA and DB at 67 and 82° C.TRRA/Si Polymerization Temp. Yield TRRA (mol/mol) (° C.) (g-PP/mg-Cat ·hr) TA 0 67 22.6 TA 7 67 12.8 TA 7 82 7.47 DB 0 67 20.8 DB 4 67 17.6 DB4 82 8.67

Table 6 shows the yield with different molar ratio of TRRA to electrondonor (TRRA/Si) for TA. The addition of TA shows almost similardeactivation degree at both temperatures where the drop in activity was−43% and −67% at 67 and 82° C., respectively, when bench marked withzero TA at 67° C. Therefore, TA is preferably used as one of thecomponents in a mixture with another TRRA. Table 6 also shows the yieldwith different molar ratio of TRRA to electron donor (TRRA/Si) for DB.The TRRA/Si ratio of 4 is for DB, a ratio that satisfies thefunctionality of reducing catalyst activity by around 60% at elevatedtemperature, 82° C.; and maintaining the catalyst activity to be notlower than 15% at normal polymerization temperature, 67° C.

A in FIG. 6 is the TREF of a polymer produced without TRRA at 67° C. Bin FIG. 6 is the TREF of a polymer produced using TA as TRRA with aTRRA/Si molar ratio of 7 at a production temperature of 67° C. C in FIG.6 is the TREF of a polymer produced using TA as TRRA with a TRRA/Simolar ratio of 4 at a production temperature of 82° C.

As can be seen from FIG. 6, the crystallinity of the polymer remainssimilar with or without using a TRRA (as exemplified by the use of TA).

Table 7 summarizes molecular weight averages and their correspondingviscosities.

TABLE 7 Molecular structure parameters obtained from SEC-IR usingconvention calibration. Pol. Temp. M_(n) M_(w) M_(z) [η] Si TRRA TRRA/Si (° C.) (×10³ g/mol) (×10³ g/mol) (×10³ g/mol) M_(w)/M_(n) M_(z)/M_(w)(dL/g) C 0 82 68 340 1000 5.1 2.9 1.73 C 0 67 85 450 1300 5.4 2.8 2.12 CTA 7 82 81 370 940 4.6 2.5 1.85 C TA 7 67 80 460 1300 5.7 2.8 2.14 C DB4 67 84 480 1400 5.7 3.0 2.19 N 0 67 72 400 1100 5.5 2.8 1.91 N PEEB 467 87 510 1500 5.9 3.0 2.30

Table 7 shows that the use of a TRRA does not significantly affect thepolymer properties. It also shows that the functionality of the TRRA isnot dependent on the type of external donor used.

Table 8 shows the yield of the polymer versus the TRRA/Si molar ratiofor PEEB with different silane donors at 67° C.

TABLE 8 Yield versus TRRA/Si molar ratio for PEEB with different Si at67° C. TRRA/Si Yield @ 67° C. Si (mol/mol) (kg-PP/g-cat) C-donor 0 24.2C-donor 4 17.9 N-donor 0 14.5 N-donor 4 8.1

In order to visualize the effect of the type of external donor used(Si), the PEEB as TRRA was tested when Si is N-donor(n-propyltrimethoxysilane) and C-donor. Table 8 shows that the TRRA,represented here by PEEB, functioning in similar pattern when the Sichanged from C-donor to N-donor. The productivity drop when TRRA addedat similar TRRA/Si ratio is the same with C and N-donors at bothpolymerization temperatures regardless of the effect of Si type oncatalyst productivity. Therefore, this example shows that regardless ofthe type of external donor used, the TRRA is effective in reducingcatalyst productivity.

A TREF was performed on the polypropylene produced using N-donor and aPEEB/Si ratio of 4 and compared to the TREF of a polypropylene producedusing the same C-donor but without using a TRRA. Also, in the case ofN-donor and PEEB, it was seen that the crystallinity of the polymerremains similar with or without using a TRRA.

CONCLUSION

A TRRA is capable of reducing the catalyst activity at a temperature of82° C. (which is the maximum desired temperature inside the expandedzone in a commercial polypropylene plant) or higher while maintainingvery good overall catalyst activity at a normal polymerizationtemperature of e.g. in the range of 65 to 72° C. (representing thetemperature of a fluidized bed in a commercial polypropylene plant).This allows a better control of the temperature in the expanded zone andconsequently the production process will be more stable.

Therefore, preferably, in the process of the invention, a streamcomprising a thermal run away reducing agent (TRRA-containing stream) isintroduced into the reactor, more preferably to the the expanded sectionduring at least part of the polymerization process. This will provide amore uniform temperature profile across the fluidized bed, leading to astable process, which in addition will lead to the production of moreuniform polyolefins.

In addition, by bringing the TRRA-containing stream into contact with atleast part of the interior surface of the expanded section,fouling/sheeting will be significantly decreased if not eliminated.

The invention claimed is:
 1. A process for the continuous preparation ofpolypropylene in a reactor from propylene and optionally ethylene and/orat least one other α-olefin monomer, wherein the reactor comprises afluidized bed, an expanded section located at or near the top of thereactor, a distribution plate located at the lower part of the reactorand an inlet for a recycle stream located under the distribution plate;wherein the process comprises: feeding a polymerization catalyst to thefluidized bed in the area above the distribution plate; feeding thepropylene and the optional at least one other α-olefin monomer to thereactor; withdrawing the polypropylene from the reactor; circulatingfluids from the top of the reactor to the bottom of the reactor, whereinthe circulating fluids are compressed using a compressor and cooledusing a heat exchanger, resulting in a cooled recycle stream comprisingliquid, and wherein the cooled recycle stream is introduced into thereactor using the inlet for the recycle stream; wherein an alkane chosenfrom iso-butane, n-butane and mixtures thereof is added to the reactor;wherein a liquid-containing stream comprising the alkane chosen fromiso-butane, n-butane, and mixtures thereof is introduced into theexpanded section during at least part of the polymerization process, isbrought into contact with at least part of an interior surface of theexpanded section; wherein the molar composition of the components in therecycle stream is chosen such that the dew temperature of the recyclestream at the reactor pressure is at least 0.10° C. below thetemperature of the reactor; and wherein the amount of iso-butane,n-butane, and mixtures thereof is chosen such that the molar compositionin the recycle stream of the total of iso-butane and/or n-butane is atleast 2 mol % and at most 5 mol %.
 2. The process according to claim 1,wherein the liquid-containing stream comprising the alkane chosen fromiso-butane, n-butane, and mixtures thereof is added to the expandedsection in a circumferential manner.
 3. The process according to claim1, wherein the liquid containing stream comprising the alkane chosenfrom iso-butane, n-butane, and mixtures thereof is added to the expandedsection via a pipe ring with a plurality of nozzles directed to theinterior surface of the expanded section.
 4. A process for thecontinuous preparation of polypropylene in a reactor from propylene andoptionally ethylene and/or at least one other a-olefin monomer, whereinthe reactor comprises a fluidized bed, an expanded section located at ornear the top of the reactor, a distribution plate located at the lowerpart of the reactor and an inlet for a recycle stream located under thedistribution plate; wherein the process comprises: feeding apolymerization catalyst to the fluidized bed in the area above thedistribution plate; feeding the propylene and the optional at least oneother a-olefin monomer to the reactor; withdrawing the polypropylenefrom the reactor; circulating fluids from the top of the reactor to thebottom of the reactor, wherein the circulating fluids are compressedusing a compressor and cooled using a heat exchanger, resulting in acooled recycle stream comprising liquid, and wherein the cooled recyclestream is introduced into the reactor using the inlet for the recyclestream; wherein a liquid-containing stream is drawn from any pointdownstream of the heat exchanger and before the inlet for the recyclestream and wherein the liquid-containing stream is introduced into theexpanded section during at least part of the polymerization process, andwherein the liquid-containing stream is brought into contact with atleast part of an interior surface of the expanded section; wherein themolar composition of the components in the recycle stream is chosen suchthat the dew temperature of the recycle stream at the reactor pressureis at least 0.10° C. below the temperature of the reactor; and whereinthe amount of iso-butane, n-butane, and mixtures thereof is chosen suchthat the molar composition in the recycle stream of the total ofiso-butane and/or n-butane is at least 2 mol % and at most 5 mol %. 5.The process according to claim 2, wherein by using at least one nozzlethat is tangential to the interior surface of the expanded sectionwherein the liquid-containing stream is added to the expanded section ina circumferential manner.
 6. The process according to claim 4, whereinan alkane chosen from iso-butane, n-butane and mixtures thereof is addedto the reactor.
 7. The process according to claim 6, wherein an alkanecomprises iso-butane.
 8. The process according to claim 6, wherein analkane comprises the liquid-containing stream is added to the expandedsection through a nozzle oriented tangential to the interior surface ofthe expanded section.
 9. The process according to claim 1, wherein analkane comprises the liquid-containing stream is added to the expandedsection through a nozzle oriented tangential to the interior surface ofthe expanded section.